Koehler. J. K., Giauque, W. F., J. Am. Chem. SOC.,8 0 , 2659 (1958). Longuet-Higgins, H. C.. Mol. Phys., 6, 445 (1963). Lunelli, B., Pecile, C.. Can. J. Chem., 46, 391 (1968). Lunelli, B., Giorgini M. G., "The Vibrational Spectrum of 1.2-Difluorobenzene and 1,2-Difluorobenzene-d4",J. Mol. Spectrosc., in press, 1976. Lunelli, B., unpublished results, 1976. Neeman, M.. Osawa, Y., J. Am. Chem. Soc., 85,232 (1963). Olah, G. A.. Halpern, Y., Shen, J., Mo, Y . K., J. Am. Chem. SOC.,93, 1251 (1971).
Olah. G. A., "Friedel-Crafts Chemistry". (a) p 126; (b) pp 331-314, Wiley-lnterscience, New York, N.Y., 1973. Scott, F. L., Oesterling, R. E., J. Am. Chem. Soc., 82, 5247 (1960). Vogel. A. I., "Practical Organic Chemistry", p 68, Longmans, Green and Co., London, 1957.
Received f o r review February 26, 1976 Accepted July 11,1976
Advances in Fischer-Tropsch Chemistry Mark E. Dry Research Department, SASOL, Sasolburg 9570, South Africa
By simultaneous manipulation of the catalyst basicity and gas composition within the reactors the product selectivity can be varied over a wide range, from 1 to over 70% methane or from zero to 50% hard wax >500 O C . Gasoline selectivity peaks at ca. 40% and diesel at ca. 20%. In low-temperature fixed-bed reactors it is the H2/CO ratio while in high-temperature fluidized reactors it is the partial pressure of COPwhich together with catalyst composition in both cases controls the selectivity. There is a fixed intercorrelation between the chain lengths of the products. The deposition of carbon on the catalyst in the fluidized beds is proportional to the value of pcoIpH: at the reactor entry and thus control of this parameter controls rate of catalyst disintegration. The o At high conversions a tail gas rate of the synthesis reaction is given by r = K(exp -€/RT) p ~ o p ~ ~ l p capH,O). of 1000 BtuISCF is produced.
+
Selectivity Spectrum Detailed information of the product selectivity spectrum in the Fischer-Tropsch reaction has been compiled over wide ranges of experimental conditions. Thus fixed-bed as well as fluidized-bed reactors have been used, temperatures have varied from 200 to 340 "C, pressures from 5 to 40 atm, and widely differing feed gas compositions have been employed. The iron catalysts have been prepared by sintering, fusion, and precipitation techniques and the influence of various promoters, both structural and chemical, have been investigated. When the selectivity results are viewed as a whole it is striking that there is a clear interrelationship between the chain lengths of the products. This is illustrated in Figures 1 and 2 where the selectivities (on a carbon atom basis) of various product cuts are plotted against either the CH4 or the hard wax (material boiling above 500 "C) selectivities. These curves signify that if a desired selectivity of any particular hydrocarbon cut is obtained, irrespective of whether it is achieved by alteration in the catalyst formulation or by adjusting the reactor process conditions, then the selectivities of all the other hydrocarbon products are effectively also fixed. The majority of the proposed mechanisms of the FischerTropsch process (Anderson, 1956; Pichler, 1963; Sternberg and Wender, 1959) assume a stepwise growth of an adsorbed hydrocarbon complex. For such stepwise mechanisms it would be expected that if the probability of chain growth was a variable then for each probability value a different product distribution would result. Figure 3 illustrates the results o f a series of such calculations. The carbon distributions were calculated from C1 to Clz0and it was assumed that the probability of chain growth of a Cs complex was always twice as large as its value for all the other surface entities. Figure 3 shows that all hydrocarbons between CH4 and hard wax exhibit selectivity peaks and that the calculated maximum selectivities for the different product cuts are in fact very 282
Ind. Eng. Chem., Prod. Res. Dev., Vol. 15, No. 4, 1976
similar to those found in practice. Thus both practice and theory indicate that in the Fischer-Tropsch process using iron catalysts it seems impossible to produce more than ca. 17%Cp or CC3, 40% gasoline, 18%diesel, and 22% medium wax (C24C:id Olefins and oxygenated compounds (e.g., alcohols and acids) are thermodynamically unstable relative to paraffins under the conditions existing in the reactors, yet they are present in appreciable quantities in the products. The olefinity of the gasoline produced can be as high as 80%and that of the Cp and C:i products up to 80 and 95%, respectively. This suggests that the olefins are primary products and the degree to which they emerge from the reactors will depend on their residence times within the catalyst particles as well as how hydrogenating the systems are. Because of this there is no general relation between say the olefinity and the carbon numbers in the way that there is between the carbon numbers themselves. For a fixed type of catalyst in a fixed type of reactor, however, there are clear correlations between these compounds and the carbon numbers. Thus Figure 4 illustrates the relation between the CH4selectivity and the selectivities of ethylene, ethyl alcohol, and acetic acid (all on a carbon atom basis) for a fluidized reactor with different feed gas compositions. Control of Selectivity It is well known from the literature that one of the key factors which controls the selectivity of iron catalysts is the amount and type of group I alkali present. The basicity of the catalyst is of course not only dependent on the alkali but also on the amount and type of other promoters or impurities present. Thus silica will react with the alkali to form less basic silicates; Le., the higher the silica content the more alkali is needed to achieve the desired basicity. Various studies have been carried out on the influence of different promoters on the surface properties of iron catalysts (Dry, 1966-1969).
I '
5t
T
t
I
30
20
10
LO
C i i L SELECTIVITY
-
50
KO
70
Figure 1. Selectivity (carbon atom basis) of various hydrocarbon cuts plotted against the CH4 selectivity.
20
10 CHI
30
SELECTIVITY
Figure 4. Carbon atom selectivity of CzH4, C~HSOHand acetic acid plotted against CHI selectivity, 30
Table I Selectivities Temp, "C
Process I
I
10
20
30
HARE W A X SELECTIVITY
-
50
10
Figure 2. Selectivity (carbon atom basis) of some of the heavier hydrocarbon cuts plotted against the hard wax selectivity.
Heavy ends production
213 227 237 247 260 288 304 310 320
Light ends production
CH4
C3
Hard wax 47 34 24 17
13 12 12
13 15
13 15 15 14 14
Table I1 Catalyst A Total Partial press, press of atm COn 29 29 29
29 29
PROBABILITY OF CHAIN GRGWTP
-
Figure 3. Selectivity of various hydrocarbon cuts plotted against the probability of chain growth. The influence of temperature is not always clear-cut and its relative effect on selectivity seems to depend on the type of catalyst, the temperature range being considered, and on the process conditions. In Table I the influence of the temperature in two different cases is compared. Thus when the process and catalyst is geared for wax production an increase in temperature markedly lowers the selectivity of the desired product while for the production of light ends the temperature is not a very important factor. The influence of gas composition is also complex in that the controlling parameters are different for different catalysts and process conditions. Thus for fixed bed wax-producing catalysts the H2/CO ratio in the reactor is the dominant factor for controlling selectivity and factors such as the partial pressure
3.78 1.52 2.00 1.38 1.52
Catalyst B
Hard H*/CO wax ratio select. 7.3 3.7 2.2
1.9 1.0
13 22
32 37 48
Total press, atm
H2/CO ratio
Hard wax select.
7.8 14.7 21.5 29.0
1.9 1.9 1.9 1.9
33 31 32 32
of COz or the total pressure are of no importance. Examples of this are given in Table 11. For fluidized bed reactors operating a t higher temperatures the Hn/ CO ratio as such seems to be of no consequence and the selectivity appears to be controlled predominantly by the partial pressure of COe in the reactor. Figure 5 illustrates the lack of correlation between CH4 selectivity and H*/CO ratio while Figure 6, where the same set of experimental results are used, shows that a strong correlation exists between the partial pressure of COz and selectivity. For these experiments the feed gas compositions, the recycle ratios, as well as the reactor pressures were varied over wide ranges but the same catalyst was used throughout. These data hence confirm previous limited evidence of the influence of the CO2 pressure on selectivity (Standard Oil, 1950). In fluidized bed reactors control of selectivity can be fully exercised by control of the catalyst's basicity and the partial pressure of Con. In Figure 7 the influence of COn on selectivity is shown for three catalysts whose basicities were low, interInd. Eng. Chem., Prod. Res. Dev., Vol. 15, No. 4, 1976
283
30
0 0 0
0 20
0
0
0
0
0
LOW 845ICITv
0
'0 -
bo
0
0
1
Figure 5. CH4 selectivity plotted aginst t h e H&O ratio of the gas a t the reactor entrance.
1
I
1
2
I
I
3
L
Figure 6. CH4 selectivity plotted against the partial pressure of CO2 a t the reactor entrance. T h e data in this figure are from t h e same set of runs used in Figure 5 .
mediate, and high, respectively. Thus by correct choice of catalyst and control of the COe pressure the selectivity can be varied at will, with CH4 selectivity ranging from 5 to 75% and gasoline from zero to 40%. From Figure 7 it will be noticed that the higher the basicity the lower the influence of the COZ pressure is.
Carbon Deposition A t higher temperatures the Fischer-Tropsch reaction is always accompanied by carbon deposition which results in swelling and break-up of the catalyst to very fine particles. In fixed bed reactors this leads to high differential pressures and eventual plugging of the reactor. In fluidized bed reactors the fine particles produced are carried out of the reactors and this loss of catalyst results in lower conversions. It is thus desirable to minimize the rate of carbon deposition for both types of reactor. The three main factors influencing carbon deposition are temperature, catalyst, and gas composition. I t has been our experience that as the basicity of the catalyst is increased the carbon deposition rate first decreases and then increases. A large number of test runs in fluidized bed reactors a t a fixed temperature and with catalysts all taken from one master batch were carried out a t different pressures, feed gas com284
Ind. Eng. Chem., Prod. Res. Dev., Vol. 15, No. 4, 1976
1
DC02 lbarl
3
1,
Figure 7. CH4 selectivity plotted against the partial pressure of COn a t the reactor entrance for three catalysts of different basicity.
1
FREE CARBON DEPOSITION PATE
I
-
Figure 8. The correlation between the factor p c o / p ~ ?(at the reactor entrance) and the rate of carbon deposition on the catalyst.
positions, and recycle ratios. From the results as a whole it emerged that the rate of carbon deposition was not related in a simple way to any individual partial pressure, e.g., of CO or of Hs, nor to the ratio HJCO. There was, however, a good correlation with the factor p c o l p ~ ?This ~ . is illustrated in Figure 8. I t follows from the nature of this parameter that, with other process parameters constant, the higher the reactor pressure the lower will the carbon deposition rate be. An additional advantage of higher pressure is that the reactor's throughput is increased in direct proportion to the pressure. Table I11 illustrates the influence of pressure. It can also be seen how with increasing pressure the CH4 selectivity decreases and that of the oil increases despite an increasing HI/CO ratio. This selectivity effect is due to the increasing partial pressure of
cos.
Kinetics The degree of conversion and the gas compositions along the length of the catalyst beds were determined in both fluidized and fixed bed reactors under various pressure, temperature, and gas composition conditions. The reactor profiles were obtained by taking gas samples a t various points in the catalyst beds as well as by taking measurements over
Table I11
Pressure, bars
Fresh feed, relative units
HdCO in reactor
9.4 12.9 14.5 18.0 21.5
1.0 1.4 1.6 2.0 2.4
4.3 5.2 5.3 5.8 5.9
pCO/pH?
Carbon deposition rate
%
CH4 select.
Cj+ select.
Conversion
29 39 43 46 47
82 85 88 88 89
pco2
x IO3
1.3 1.7
-
-
6.7
1.8
5.7
2.2 2.3
3.1
3.3 2.1 1.6
26 19 17 15
1.7
1.1
11
OCO
=
PCO
+ aPHzO where a = kHnO/kCO. The rate equation thus becomes
,.=
PCO
PCo'PHz
k(e-E/RT)
PCO
+ QPHzO
The value of the activation energy ( E ) is 6 for the hightemperature fluidized bed and 15 kcal/mol for the low-temperature fixed-bed processes. While from the above equation it appears that COz plays no direct role in the reaction rate, it does play an important indirect role via the water-gas shift reaction which occurs readily under the process conditions. I BED LENjTH
Hz0
-
Figure 9. Plot of the accumulative conversion (moles converted per hour) against the catalyst bed length for the fluidized and fixed-bed reactors. catalyst beds of varying lengths. Typical profiles are illustrated in Figure 9. A simple kinetic equation derived from first principles was found to fit the curves satisfactorily. If the slow step in the Fischer-Tropsch synthesis is the reaction of a molecule of H2 with a chemisorbed CO molecule, namely
+ CO
-
COz + H2
In the fluidized bed the shift reaction is at equilibrium all along the length of the bed while in the lower temperature fixed bed the simple equation r = mpco adequately accounts for the rate of this reaction. In a differential reactor where the concentration of products and hence of water will be very low and so can be neglected, the Fischer-Tropsch rate equation will simplify to
r = KPHr which is in agreement with experience in such reactors (Dry, 1972).
High Btu Gas Production
which in many reaction schemes is considered as the initiating reaction, then the rate of reaction will be
Using Langmuir's adsorption theory, the fractional CO coverage of the active catalyst surface in a system in which CO is competing with H20, CO2, and H2 for the adsorption sites will be given by BCO
=
kcoPco 1 + kcopco
+ ~ H ~ O P H+~kcogco2 O +~
H $ H ~
Based on knowledge of the relative strengths of adsorption of the various gases on iron catalysts (Dry et al., 19691, the above equation can be simplified to
The tail gases from Fischer-Tropsch reactors, irrespective of whether they be wax, gasoline, or light ends producers, can, after knock-out of the water, be further processed in hightemperature fluidized bed reactors. These tail gases, after scrubbing, have heating values of 800 to 1000 Btu/SCF depending on the degree of conversion. It is not easy to produce the high BTU gas in a single reactor as the water vapor product inhibits the reaction rate over iron catalysts and thus makes the attainment of the necessary high degree of conversion difficult. The two-stage process, however, with intermediate water knock-out makes this possible. Table IV illustrates a few examples. In spite of the relatively high concentrations of Hz in these gases their heat contents are high because of the Cp and C3 hydrocarbons also present. If the and CB hydrocarbons in the primary reactor tail gases are desirable products an alternative scheme is to extract them and then process the remaining gas over a nickel catalyst thus producing a gas, after C 0 2 scrubbing, consisting mainly of CH4.
Table IV Primary reactor selectivity, %
Secondary reactor Vol % of scrubbed tail gas
CHs
Ci+
H?
Inerts
CO
CHI
C?
C?
BtuISCF
47 36 9
11 18
19.2 16.9 16.1
3.4 3.3 5.1
0.2 3.4 0.6
66.9 65.4 68.4
7.0 6.6 5.3
3.3 4.4 4.4
980 1000 970
54
Ind. Eng. Chem., Prod. Res. Dev., Vol. 15, No. 4, 1976
285
Nomenclature
BCO = surface coverage by chemisorbed CO etc. = partial pressures of Hz, CO etc. hco = ratio of rates of adsorbtion to desorption E = activation energy
P H ~ p, c o
Dry, M. E., Shingles, T., Boshoff, L. J., J. Catal., 25, 99 (1972). Pichler, H., Brennsfoff-Chem., 44, 33 (1963). Standard Oil, British Patent 647 079 (Dec 6, 1950). Sternberg, H. W., Wender, I., "International Conference on Co-ordination Chemistry", p 35,Chemical Society, London, 1959.
Received for review March 23,1976 Accepted June 22,1976
Literature Cited Anderson. R. B., "Catalysis," Vol. 4, p 123, Reinhold. New York, N.Y.. 1956. Dry, M. E., Brennstoff Chem., 50, 193 (1969). Dry, M. E., et at.. J. Catal., 6, 194 (1966); 7, 352 (1967); 11, 18 (1968). Dry, M. E., Shingles, T., Boshoff. L. J., Oosthuizen, G. J., J. Catal., 15, 190 (1969).
Presented at the Division of Industrial and Engineering Chemistry, Centennial Meeting, American Chemical Society, New York, N.Y., April 1976.
The Ion-Exchange Membrane, Chlor-Alkali Process Maomi Seko Asahi Chemical lndustry Co., Ltd., Yurakucho,Chiyoda-ku, Tokyo, Japan
The first commercial membrane chlor-alkali plant with an annual production capacity of 40 000 metric tons of caustic soda has been running successfully since April 1975. Operation results and the requirements of membranes are described in detail. A new perfluorocarboxylic acid membrane is compared with a perfluorosulfonic acid membrane. A perfluorocarboxylic acid membrane having a high current efficiency even at a high concentration of caustic soda is distinguished from a conventional membrane. A unique design is described for the metal electrolyzer and heat recovery evaporator for its membrane chlor-alkali process. Compared with conventional chlor-alkali processes, the membrane process is a better energy saving method. It requires lower costs for investment, raw materials, and utilities.
Introduction Since 1968, Asahi Chemical has been developing a chloralkali process using an ion-exchange membrane. First research efforts were directed toward the development of a threecompartment method using a hydrocarbon ion-exchange membrane and a porous diaphragm. Since 1974 efforts were concentrated on the development of a two-compartment method using ion-exchange membranes of the perfluorosulfonic acid type. In 1973, in order to solve the mercury pollution problem, the Japanese Government decided to convert all chlor-alkali plants using mercury processes to nonmercury processes. I t was decided that as the first conversion step, two-thirds of the mercury process plants (approximately 3 000 000 metric tons per year of caustic soda) would be converted to nonmercury processes by the end of September 1975 and secondly, the remaining mercury process plants would be converted by the end of March 1978. In compliance with this policy, Asahi Chemical constructed a membrane chlor-alkali plant with an annual production capacity of 40 000 metric tons as 100%concentration caustic soda and started its commercial operation in April 1975. This plant has been running quite satisfactorily since its start-up and has produced caustic soda of 17.6 to 21% concentration. The electrolyzer and cell compartments employed in this plant have various advantages. In addition, the electrolyzer was installed in conjunction with an evaporation system of unique design. In the meantime, a perfluorocarboxylic acid type membrane of excellent electrochemical performance was developed. The evaporation system is characterized by the heat recovery evaporation system which requires only a very small amount of energy when caustic soda of 25 to 30% concentration is concentrated to 48% concentration. 286
Ind. Eng. Chem., Prod. Res. Dev., Vol. 15, No. 4, 1976
Principle of Electrolysis As shown in Figure 1, when a salt solution is electrolyzed while brine is supplied to the anode compartment partitioned by a cation-exchange membrane, Na+ ion migrates through the cation-exchange membrane to the cathode compartment by an electric force. On the other hand, transfer of C1- to the cathode compartment for the anode compartment is prevented by the cation-exchange membrane. Thus high-purity caustic soda solution is obtained in the cathode compartment. Figure 1shows the material balance and electrolysis products. (1 - 7 ) equivalent of OH- ion (7 = current efficiency) migrates from catholyte to anolyte under 1Faraday conduction of electricity. In order to neutralize this OH- ion, (1- 7 ) equivalent of hydrochloric acid is injected into the anolyte stream. Water is injected into the catholyte stream to maintain the concentration of caustic soda obtained in the catholyte. Chlorine gas is generated in the anode compartment. The hydrocarbon-type cation-exchange membrane is attacked by Clz or C10- ion and scission of the polymer chain in the membrane occurs in a short time. Electrolysis terminates in a short period of time. Since the perfluoro sulfonic acid membrane Nafion (Connoly and Gresham, 1966) was developed by E. I. du Pont in 1960's, Nafion was the sole membrane applicable to the chlor-alkali cell because it does not contain C-H bonds. However, membranes having sulfonic acid exchange groups cannot have high current efficiency in strong concentrations of caustic soda due to the counter migration of OH- ion from catholyte to anolyte. Figure 2 shows the relation between the catholyte concentration and current efficiency of various types of Nafion membranes. Nafion 315 in the form of a laminated sheet of
